Catalytic conversion of mixtures of alkyl chlorides and petroleum fractions



Aug. 19,1947. H. T. BROWN Erm.

CATALYTIG CONVERSION OF MIXTURES OF ALKYL CHLORIDES AND PETROLEUM FRACTIONS Filed Dec. 26, 1944 Patented` ug. 19, 1947 cATALYrrc CONVERSION oF MIXTURES oF ALKYL onLoRIDEs AND PETROLEUM FRACTIONS Henry Trueheart Brown and Everett Gorin,

s Dallas, Tex., assignors, by mesne assignments, to Socony-Vacuum Oil Company, Incorporated, New York, N. Y., a corporationoi' New York Application December 26, 1944, Serial No. 569,837

' (c1. aso-,666)

Y 6 Claims.

'Ihls invention relates to th'e conversion of normally gaseous hydrocarbons to hydrocarbons of v higher molecular weight. `More particularlyjthis invention relates to the conversion of methane and natural gas to a mixture of normally liquid hydrocarbons of high octane number for use as mbtor fuel by converting the gaseous hydrocar bons tothe corresponding halides and condensing the halides by means of a catalytic dehydrohaloi gene-condensation process. Specifically tlie present invention has todo with the catalytic dehydrohalogeno-condensation of alkyl halides having less than four carbon atoms to the molecule, such as the chlorides'of methane, ethane and propane in a reaction zone wherein petroleum fractions boiling above the gasoline range are undergoing simultaneous conversion to produce hydrocarbon fractions suitable for incorporation in motor fuel."

It is known to condense light alkyl halides by pyrolysis. For example, in the .patent obtained by one of the presentinventors, U. S. 2,320,274,

entitled Conversion of normally gaseous hydrocarbons, it is taught to halogenate the hydro- .carbons and then to pyrolize the light alkyl -halides at temperatures above 500 C. to produce benzene, acetylene and ethylene. In the cepending application of Everett Gorin and Manuel I-I.l Gorin entitled Catalytic conversion 'of normally gaseous hydrocarbons, Serial No. 556,746, filed October 2, 1944, now abandoned and reiiled as Serial No. 677,402, on June 17, 1946, as a continuation-.in-part, it has also been taught and claimed to catalytically dehydrohalogenate and condense methyl chloride to hydrocarbons having more than one carbon atom per molecule at temperatures below 500 C. We have found it advantageous to carry out the catalytic dehydrohalogenation-condensation of alkyl chlorides such as methyl chloride in the presence of hydrocarbons such as. petroleumnaphthas or a petroleum gas oil cut or in the presence yof heavy synthetic oils such as a diesel fuel or gas oil cut from the Fischer-Tropsch' product. -I-lighly refractory gas` oils containing polyalkylated aromatics such as hexamethyl benzene and polycyclic aromatics suchas the alkyll naphthalenes may also be added although these materials are cracked to a much less extent than a parainic type of gas oil. By

the term naphthas, as used in the specification and claims, wemean natural gasoline distillate stocks such as casinghead gasoline, light and heavy naphthas obtained from petroleum, and distilling within the range of'fromabout '10 F. to about 500 F. By the term gas oil, we mean that' fraction of natural petroleum or mixture of synthetic hydrocarbons the distillation temperamethane and natural gas, hydrocarbons having molecular Weights higher than the majorhydrocarbon components of natural gas which hydrocarbons in turn'are susceptible to conversion to motor` fuel. A further object is to catalytically condense the halide derivatives of methane and the halide derivatives of other hydrocarbon components of natural gas to higher molecular weight hydrocarbons in the same reaction zone wherein a hydrocarbon or normally liquid mixture of hydrocarbons of boiling range in or above that of motor gasoline is simultaneously converted to a mixture of hydrocarbons suitable for incorporation in motor fuel. In particular, itis the object of our invention to suppress the `formation, of .carbon and methane in the dehydrohalogenation of methyl chloride by carrying out the reaction in the presence of a high boiling hydrocarbon oil. These and other objects will be apparent from th'e description of our invention.

The dehydrohalogeno-condensation of methyl chloride is eiIected by the aid of catalysts which possess both dehydrohalogenation and polymerization activity. Suitable catalysts are those which comprise an association of amphoteric and acidic oxides such as alumina on silica, zinc oxide on silica, boric oxide on alumina,`beryl liumoxide on silica or gallium oxide on silica. Other mixed oxide catalysts such as magnesiasilica, zirconia-silica or thoria-silica catalysts may be used. Pure amphbteric or acidic oxides may be used alone, such as for example, alumina, silica or titania. Active catalysts may also Ibe prepared from naturally occurring materials such as by acid treating clays of the fullers earth, bentonite, montmorillonite or attapulgus variety. These materials all show activity in the cracking of petroleum fractions and hence are suitable for use in our process. However, gel type catalysts having high porosity are preferred, and it is highly desirable that the catalyst be capable of withstanding high temperatures such as th'ose produced by oxidative regeneration since in the reaction cycle considerable carbon is deposited on the catalyst which must be removed in order to restore the catalyst activity. Alumina-silica catalysts prepared according to the procedure described in the above Serial No. 677,402 and containing from one to twenty percent alumina are particularly adaptable to use in our process.

We prefer to carry out the reaction under conditions such that a relatively large proportion of the gasoline product is derived from the methyl chloride. Hence a high ratio of methyl chloride to oil is employed. The weight percent of methyl chloride in the feed should be greater than weight percent and preferably greater than 30 weight percent, while amounts as high as 85 or 90 weight percent of methyl chloride may be employed. The above catalysts are comparatively inactive at temperatures below 250 C. while in y most cases temperatures above 300 C. are required to affect a reasonably rapid reaction. Tem.. perature of operation should in no case exceed 500 C. and we prefer to operate at temperatures below 450 C., that is, Within the range of from about 300 C. to about 450 C. l

A wide range of pressures may be employed in converting the mixture of methyl chloride and heavy hydrocarbon oil to motor fuel blending products, such pressures varying from atmospheric to as high as 1800 or 2000 pounds per square inch. When the process is carried out in the neighborhood of atmospheric pressure and in the upper temperature range, both liquid and gaseous products are highly unsaturated. If pressures of 100 to 500 pounds per square inch are used, the temperature of operation being held below 425 C., the liquid products are substantially saturated while the oleiin content of the gaseous products is relatively low. We prefer to operate at pressures from about atmospheric to about 500 pounds per square inch. The space velocity is adjusted to convert from 50 to 80 percent oi' the methyl chloride per pass and may vary within the range of from about 0.2 to 10.0 volumes of total liquid feed, that is, oil plus methyl chloride, per volume of catalyst space per hour depending on the temperature and pressure employed.

Relatively large quantities of isobutane are`produced when operating under the above condi tions which may be advantageously recycled to the process where it is converted by condensation with the methyl chloride to hydrocarbons of lower volatility suitable for incorporation in motor fuel. The isobutane also acts as a diluent for the methyl chloride and aids in keeping the instantaneous concentration of the methyl chloride at a low figure which is desirable for the most successful operation of our process since the rate of carbon formation is greatly reduced thereby. We may also maintain a low instantaneous concentration of methyl chloride by introducing the vaporized feed at a multiplicity of points to the stationary catalyst bed described below in our preferred method of operation.

Referring now to the drawing, towers I and 2 contain alumina gel-silica gel catalyst either as a continuous bed or disposed in trays. We prefer the latter method of packing these towers since the diiiiculty of controlling regeneration temperatures is reduced if the catalyst is maintained in a series of relatively shallow beds. Towers I and 2 are manifolded in such a manner that while one of these is on stream for the simultaneous condensation and cracking reactions, the other is undergoing oxidative regeneration thus providing substantially continuous operation.

A feed comprising 50 to 75 parts by weight of methyl chloride admixed with a petroleum naph tha. or gas oil is passed by means of pump I0 in line II to heat exchanger I2 at a pressure of `from about 250 to about 300 pounds per square inch where heat is absorbed from the reaction product in line I3 and then flows throughA line I4 and heat exchanger li'where additional heat is absorbed from hot regeneration gases in line I6. From exchanger I5 the hot feed passes via line I 1 to manifold line I8 provided with valves |19 and 20 and connected with reactor feed manifold lines 2l and 22. With valve I9 in line I8 open and valve 20 closed, the feed is introduced to tower I through a multiplicity of valved feed lines leading from manifold line 2I at a temperature of about 400 C. and at a space velocity in tower I of about 3 volumes of liquid feed per volume of catalyst space per hour. Isobutane and ethyl chloride product described hereinbelow may be recycled by means of pump 23 in line 30 which joins manifold line 3l provided with valve 32 which is open and valve 33 which is closed. The isobutane and ethyl chloride product maybe introduced to tower I via the main feed stream by closing valve 34 in line 30 and passing this recycle product through line 35 to recycle line |03 which joins line II. Additional isobutane from the product is also furnished in the form of recycle methyl chloride-isobutane azeotrope, as hereinafter described, via recycle line |03.

The product from reaction tower I consisting of methane, ethylene, ethane, hydrogen chloride, unreacted methyl chloride, ethyl chloride, propylene, propane and C4 hydrocarbons along with higher boiling hydrocarbons and unconverted gas oil is passed by drawoff line 40 to manifold product line 4I which is provided with valvesl 42 and 43, the latter being closed and the former open for delivery of the product to line I3 leading to heat exchanger I2 where the product is subjected to initial cooling. From exchanger I2 the product passes via line 44 to cooler 45 wherein the temperature of the product is further reduced to cause condensation of isobutane and higher boiling components of the decomposition product and adsorption of these in the gas oil. A considerable part of the (h hydrocarbons and of the hydrogen chloride decomposition product will also be absorbed in the liquid gas oil. From cooler 45 the relatively cold product passes via. line 46 to vapor release drum 4l which is operated at approximately the same pressure as reaction tower I. Non-condensible gases comprising methane, C2 hydrocarbons and a part of the C: hydrocarbons and HC1. pass overhead from drum 4l through pressure release valve 48 in line 49 to the hydrocarbon chlorination process for the production of alkyl chlorides as described and claimed in the copending joint application of C. M. Fontana and E. Gorin, entitled Manufacture of halogenated hydrocarbons, Serial No. 548,351, led August 7, 1944. These alkyl chlorides may then be used as feed to our process.

The liquid product from drum 41 is passed by means of pump 50 in line 5I to fractionator 52 provided with heating means 53. Fractionator 52 operates as a stripper for removal of residual noncondensible hydrocarbon gases and hydrogen chloride from the condensible hydrocarbons and unconverted gas oil product. The pressure in tower 52 is maintained at 400 to 500 pounds per square inch. The overhead gas passes via line 54 to condenser 55 and the condensate passes thence via line 56 to redux drum 61. Uncondensedgas` comprising the last traces of methane and the Cz gases pass overhead from drum 51 the pressurey being controlled by means of valve 58 in line 69 which connects withline I9 leading to the alkyl chloride manufacturing. process. VHydrogen chloride" condensate in reux drum 51 is removed therefrom by pump 60 in line 6 I` to be returned in ducing alkyl chloride feed stock but `also desirable to remove the hydrogen chloride from therecycle stock since large-amounts of hydrogenchloride have an inhibiting eifect on the course of the ferred via line 90 and pressure .valve 9| to deisobutanizer tower 92 which is provided with i reboiler means 93.. Methyl chloride forms an azeotrope4 with isobutane and hence the overhead i from tower 92 consists primarily of methyl chloride-isobutane azeotrope boiling at about 25.6 C. at atmospheric pressure, provided these components `are present in the proper proportions. This azeotrope consists f approximately 7'1 mol percent methyl chloride and 23 `mol percent isobutane. Provision is made for adding methyl chloride to line 90 via line 94 by means of pump 95 for increasing the amount of methyl chloride in tower 92 when the conversion of methyl chloalkyl chloride condensation reaction. It is desirable to maintain the vhydrogenchlorlde con-` centration below 35 percent of the total gas and in any case the concentration should not constitute 9ver`60 mole percent of the total conversion products in reactor eiiiuent line I3.

The liquid product from fractionator 52 is transferred via pressure release valve 64 in `line 65 to debutanizer tower 66 for removal of propane, propylene, the C4 fraction, unconverted metlwl chloride and ethyl chloride from the higher boiling product. The overhead from tower 66 may also contain hydrogen chloride in minor amounts which is retained in the liquid product from tower 52 as a result of the relatively high pressure fractionation in tower 52. Tower 66 is operated at a pressure of from about 100 to about 200 pounds per square inch and is provided with reboiler means 61. The overhead gaseous product from tower 66 passes via line 68, condenser 69 and line 10 to reflux drum 1| from which noncondensed gas, primarily hydrogen chloride, passes via valved line 12 to line 63 for use in producing methyl chloride or other alkyl chlorides as described in the aforementioned Serial No. 548,351. Condensate from reflux drum 1| is picked up by pump 13 in line 14,V 9,. part being transferred through valved line 15to tower 66 as reflux and the remainder is passed via valved line 16 to depropanizer tower 11. 'I'he components of the stream in line 'I6 are all suitable for recycle to the reaction zone and hence valved line 30 is provided,which` leads from line 16 directly to manifoldline 3|. As stated hereinabove, this recycle product may be diverted to the main feed stream in line I`I via valved line 35 and recycle i line |03.

That part of the stream of relatively light hydrocarbons in line 16 not diverted as recycle through line 30 is fractionated in tower 11 to separate propane and propylene from the C4 hydrocarbons and alkyl chlorides. Tower 11 is operated at a pressure of from about 200 to 250` pounds per square inch with tower top temperature maintained at about 35 C. to about 45 C.; tower 11 being provided with reboiler means 16. The overhead consisting of propylene and propane passes via line 19, condenser 80 and line 8| toy reflux drum 92 whence any` uncondensed propylene passes overhead through line 83 to recycle line 30. Condensate is removed from drum 82 by means of pump 84 in line 85 for transfer in part to tower 'Il as reflux through line 86 and the remainder, comprising relatively pure propane, may he sent to storage v ia line 81 or it may be recycled via line 30 to reaction tower I.

The bottom product from tower 11 is transride in reaction tower I is isumciently4 high to leave insuiicient unconverted methyl chloride to l form the desired azeotrope with isobutane product. The overhead product from tower 92 consisting of the azeotrope and a relatively small amount of propane passes overhead through line 96, condenser 91 and line 98. to reflux drum 99 whence the condensate is sent in part by means of pump |00 in line I0| and via valved line |02 to tower 92 as reflux, fthe remainder being passed via valved line |03 as recycle to main feed line Il. Any accumulation of gaseous propane in drum 99 is removed via line |04 to storage. The C4 bottom fraction .which is removed from tower 92 via. line |05 contains s ome ethyl chloride which maybe separated by scrubbing with an alcoholwater mixture andthe separated ethyl chloride may -be recycled to reaction tower I whil'e the C4 fraction which may-contain excess isobuta'ne over that required to form the azeotrope with methyl chloride may be withdrawn from the system and y subjected to alkylation and/or isomerization re'Y actions to produce aviation grade gasoline.

Returning now toldebutanizer tower 66 the bot- ||4 in line I|5, `a partpf the liquidbeing returned totower |08 as reflux through line |I6 and the remainder is withdrawn through line I I1 and given la caustic Wash to remove any traces of hydrogen chloride -prior to incorporation in motor fuel. The bottom product from tower |08 of distillation range above that of' motor gasoline and consisting primarily of unconverted gas oil is Withdrawn through line I I8 and is recycled via line |03 to feed line |I a part being eliminated from the process through line |I9 to prevent excessive accumulation of highly cyclized refractory cycle stock.

When a reactor has been on stream for an interval up to minutes depending on the operating conditions relative to space velocity, temperature and pressure, suflcient carbon will have accumulated on the catalyst` to necessitate reactivation. This is accomplished after purging vthe `reactor of hydrogen chloride and other vapors by passing air through the reactor, the air being diluted with flue gas in order lto control the rate of oxidation and to maintain the regeneration temperature below about 650 C., preferably below about 600 C. It is desirable to purge the reactor of hydrogen chloride and hydrocarbon vapors before regeneration since the combustion of the hydrocarbons results in the formation of water vapor which in the presence of hydrogen chloride aggravates corrosion problems. The purging and reactivation of the catalyst in reaction zone 2, that is tower 2, is carried out while tower is on stream as described hereinabove.

Tower 2 is isolated for purging before regeneration by closing valve |23 in line |22, valve 33 in line 3|, valve 20 in line I6, valve |43 in line |42, valve |45 in line I 46, valve'43 in line 4| and valve |25 in line 40. Methane purge gas is introduced via line |2| and passes via lines |22 and 3| .to reactor 2 for removal of residual hydrogen chloride, unreacted alkyl halides and hydrocarbon vapors. The purge gas leaves reactor 2 through line I5| and thence passesv through open valve |26 in manifold line |21 and through open valve |29 in line |30 to line 49 for transfer to the alkyl chloride production unit described in the hereinabove mentioned copending application, Serial No. 548,351. The methane purging step is followed by a short flue gas purge to clear reactor 2 of methane prior to the reactivation step.

Referring now to the reactivation step following the purgingof reactor 2, a mixture of air and flue gas is introduced to the system by means of compressor |40 in line 4|. The extent of dilution of oxygen in the regenerating gas should be such that temperatures in excess of 650 C. are avoided and prefera'bly the temperature should not exceed 600 C. since synthetic alumina-silica catalyst and also acid treated clays tend to become permanently deactivated at higher temperatures. I

The flue gas diluted air is passed from line I4| to manifold line |42, which is provided with valves |43 and |44, and thence to line 22 for introduction to tower `2 via manifold and valved feed lines leading from line 22. The vregeneration zone is isolated by closing valve |44 in line |42, valve 20 in line I8, valve I4`| in line |46, valve 43 in line 4| and valve |26 in line |21 and provision for continuous ow of regeneration gas is made by maintaining valve |43 in line |42 and valve |45 in line |46 in the open position. The regeneration gas passes from tower 2 through line I5| to line |46 and thence through line I6 to heat exchanger I5 whence it is eliminated from the system through valved line |49. If desired, a part of the spent regeneration gas may be recycled to line |4| through valvcd line |50 for dilution of fresh regeneration air. The flow of regeneration gas is continued for a temperature adjustment period following complete reactivation of the catalyst until the catalyst bed temperature has lowered to such a level that fresh feed Awill be held within the desired temperature range for the conversion cycle. The purge gas should contain no oxygen during this purging operation and hence steam may be substituted for the flue gas for the nal temperature adjustment operation.

Although we have described our process as being carried out at pressures of 250 to 300 pounds per square inch, higher or lower pressures may be used. The simultaneous condensation and catalytic cracking reactions may be carried out at lower pressures, for example, of the order of atmospheric to 50 pounds per square inch gauge. However, such type operation involves greater expenditure of capital for compression equipment for separation of predominantly gaseous hydrocarbon product. Certain advantages favor low pressure operation. For example, the carbon deposited on the catalyst may be removed more advantageously by operating the oxidative reactivation step by the well known hindered flow type procedure which is adaptable only to relatively low pressure operation. Oxidative regeneration temperatures are more uniform and more easily controlled in this type operation. Hindered flow, that is, suspended powdered catalyst operation may also be used for the conversion cycle if the pressure is maintained below about 'Z5 pounds per square inch. We may also adopt other methods of vcatalytic contacting to our process of simultaneously converting alkyl chlorides and naphthas or gas oil to liquid hydrocarbons of motor fuel distillation range. For example, the various types of moving catalyst bed techniques well known in the art of hydrocarbon conversion, particularly in the art of converting petroleum hydrocarbons and hydrocarbon mixtures, may be used if relatively low pressure operation is followed.

The improvements effected in the dehydrohalogenation-condensation of methyl chloride and in the cracking of cetane by carrying out the processes simultaneously are illustrated by the experiments listed in the table below. All the experiments were carried out at 425 C. and atmospheric` pressure with fresh samples of an alumina-silica mixed gel catalyst prepared according to the procedure outlined in the above copending application Serial No. 556,746. The mol ratio of silica to alumina in the catalyst was about '1/ 1.

Experimental NumbeL l 2 3 4 1 Reacumts CMH cmcl..{jg; g'll j j}cunn Vapor Space Veloc- 2.79. 2.73 2.85 0.65.

ity in.. Percent CuHu Dc- 24.7. 54.0 54.0.

composed. Percent CI'IaCl De- 57.8 68.4

composed. Wt. Percent Carbon 4.2.. 7.8 7.8.

u Wt. Percent C114 5.3 5.5 5.5.

(CuHu). M(o Percent Carbon 24.7 12.2

3 Mol Percent CHO 10.2 2.9

(CHaCl). Yield 01+ (CHCI).. 65.1. 84.9

l Percent by weight.

The comparison of the behavior of mixtures of cetane and methyl chloride with the pure materials by themselves was made at the same value of the vapor space velocity, i. e., at approximately the same times of contact with the catalyst. The space velocity is expressed in terms of volume of vaporous feed, calculated as the volume the material would have if it were a perfect gas at normal conditions of temperature and pressure. passed over unit volume of catalyst per unit time.

The weight percent yields of carbon and of methane are calculated from the weight of carbon and methane formed respectively, divided by the total weight of C1 through Ca hydrocarbons plus carbon formed in the cracking of the cetane. The yields of carbon and methanelderived from the methyl chloride are expressed in terms of mols The advantages of our process are seen quite clearly from these examples. Comparing experiments 1, 2, and 3, it is seen that by carrying out the dehydrohalogenation of the methyl chloride I dividually are reduced from 24.7 and 10.2 percent j to 12.2 and. 2.9 percent, respectively.

Having thus described the nature' of our invention and means for practicing the same, but without 'intending thereby to limit our invention to Y such specicmeans, save as expressly set forth in the appended claims, we claim as new:

1. A process for the condensation of alkyl chlorides having less than four carbon atoms per molecule 4to produce hydrocarbons having a greaternumber of carbon atoms than the alkyl chloride treated which comprises passing a stream consisting essentially of a mixture of said alkyl chloride and a higher than gasoline-boiling petroleum fraction containing from' 15 to 90 weight per cent of said alkyl chloride through a reaction zone containing a, solid, metallic oxide hydrocar boncracking catalyst at a temperature of from about `300" C. to about 450 C. and at a rate such that from about 50% to about 80% of the alkyl chloride content of the feed mixture isconverted per pass, and recovering the hydrocarbon content of the eilluent stream from the reaction zone.

. 2. A process for the-condensation of alkyl chlorides having less than four carbon atoms per molecule to produce ,hydrocarbons having a greater number of carbon atoms'than the alkyl chloride treated which comprises passing a stream consisting essentially of amixture of said alkyl chloride and a gas oil containing from V154 to 90 weight per 'cent of said. alkyl chloride through a reaction zone containing a solid, metallic oxide hydrocarbonf'cracking catalyst at a temperature of from about 250 C. to about 500 C, and at a' rate such that from about 50% to about carbon contento! the eilluent streamfrom the reaction zone.

4. A process for the condensation of methyl chloride to produce hydrocarbons or at least two carbon atoms therefrom which comprises passing a stream consisting essentially of a mixture of said methyl chloride and a. gas oil and containing from 30 to 90 weight per cent of said methyl chloride through a reaction zone containing a solid, metallic oxide hydrocarbon-cracking catalyst at a temperature of from about 300 C. to about 450 C. and at a rate such that from.

`about 50% to about 80% ofthe methyl chloride contentof the feed mixture is converted per pass, and recovering the'hydrocarbon content of the eilluent stream from the reaction zone.

5. A process for the condensation of alkyl chlorideshaving less than four carbon atoms per .mole'cule to produce hydrocarbons having a greater number of carbon atoms than the alkyl chloride treated which comprises passing a lstream consisting esentially of a mixture of said alkyl chloride and a higher than gasoline-boiling petroleum fraction containing from 15 to 90 weight percent of said alkyl chloride through a reaction zone containing an alumina-silica catalyst at a temperature of from about 300 C. to about 450 C. and at a, rate such that from about 50% to about 80% of the alkyl chloride content of the feed mixture is converted per pass, and recovering the hydrocarbon content of the eiiluent stream from the reaction zone.

6. A process for the condensation of methyl \4 chloride to produce hydrocarbons having at least two carbon atoms per molecule therefrom which comprises passing a stream consisting essentially of a mixture of said methyl chloride and a gas oil and containing from to 9.0 weight per cent of said methyl chloride through a reaction zone containing an alumina-silica catalyst at a temperature of from about 300 C. to about 450 C. and at a rate such that from about to about 80% of the methyl chloride content of the feed 80%of the alkyl chloride content of the feed mix-V I ture is converted per pass, and recovering the hydrocarbon content of the eiiluentstream from the reaction zone. v

3. A processjor the condensation of alkyl chlochloride treated which comprises passing a stream consisting essentially of a mixture of -said alkyl chloride and aA higher than gasoline-boiling petroleum fraction containing from 30 to 90 weight `per cent of said alkyl chloride through a reac tion zone containing a solid, metallic oxide hydrocarbon-cracking catalyst at a temperature of from about 300" C. to about 450 C. and ata .rate such that from about 50% to about 80% of the alkyl chloride content of the feed mixture is converted per p ass, and recovering the hydromixture is converted per pass, and recovering the hydrocarbon `content of the eiiluent stream from the reaction zone.

HENRY TRUEHEART BROWN. EVERETT GORIN.

REFERENCES CITED UNITED STATES PATENTS Date Number Name 1,878,262 Chappell Sept. 20, 1932 2,201,306 Subkow May 21, 1940 `2,213,345 Marschner Sept. 3, 1940 2,220,090 Evering Nov. 5, 1940 2,350,159 Folkins May 30, 1944 2,368,446 Buc Jan. 30, 1945 v 1,826,787 Howlett Oct. 13, 1931 FOREIGN PATENTS n Number Country Date 513,674 Great Britain Oct. 19, 1939 OTHER REFERENCES i Lu et al., article in The Renner, vol. 20, No. 9; Sept. 1941; pages 13D-133, 26o-683.

Certificate of Correction Patent No. 2,425,861. l August 19, 19.17;-

HENRY TRUEHEART BROWN ET AL.

It is hereby certified that error appears in the printed specification of the above numbered patent re uiring correction as follows: Column 8, line 42, first column of the table, for"Mol ercent CHO read Mol Percent CH4; and that the said Letters Patent should be read with this correction therein that the same may conform to the record of the case in the Patent Oiice.

Signed and sealed this 2d dey of December, A. D. 1947.

THOMAS F. MURPHY,

Assistant Oommz'asoner af Patents. 

